Gas Pre Treatment

Gas Pre-Treatment and their Impact on Liquefaction Processes J.M. Klinkenbijl M.L.Dillon E.C. Heyman Shell International Oil Products Research & Technology Centre Amsterdam

Presented at GPA Nashville TE meeting 2nd March 1999 Summary Natural gas generally requires removal of H2S, CO2, COS, organic sulphur compounds, mercury and water prior to liquefaction in order to meet product specifications, avoid blockages and to prevent damage to process equipment. The cost of pre-treatment is dependent on the type and concentrations of the contaminants in the natural gas. Most of the operational base load LNG plants process feed gases with only low concentrations of CO2, mercury and water as contaminants. This type of gas requires the ‘minimum’ of treating, often comprising of a CO2 removal unit, molecular sieves for drying and a carbon bed for mercury removal. The Shell Sulfinol process is the most widely applied acid gas removal process, serving some 40% of the installed base load LNG capacity, and has proven to be very reliable and cost effective. When mercaptans are present in gas feed, the Shell Sulfinol process is strongly preferred as the acid gas removal step, since it combines total CO2 and H2S removal with mercaptan removal up to 97%. Formulated MDEA solvents have a comparable capital cost to Sulfinol, but lack the mercaptan removal capabilities, with one exception being the Flexsorb formulation (from Exxon) also containing sulfolane. Revamp of a gas pre-treatment unit from limited mercaptan handling capability to significant mercaptan handling capability can elegantly be done using an integrated Sulfinol based concept. The relative capital investment for acid gas removal in a LNG plant increases significantly with increasing CO2 content. At 2 mol% CO2 the acid gas unit represents 6% of the processing equipment cost but at 14 mol% CO2 it represents 15% of the processing equipment cost. The capital cost for dehydration and mercury removal depend mainly on the natural gas throughput. New developments such as membrane technologies are starting to be considered as an option for bulk removal of CO2 but solvent absorption remains the only cost effective treatment process for meeting LNG specifications. Further developments may change this in the future.

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1. Introduction The type of contaminants and their concentration in the natural gas affect the overall LNG production cost. In the past most feed gases for base load LNG plants only contained CO2 (<< 10 mol%) with traces of H2S. This type of gas requires the ‘minimum’ of pre-treatment. As the market for LNG has expanded and more gas fields become economically viable to develop there has been a need to treat feed gases richer in CO2 , H2S, COS and mercaptans. This naturally increases the cost of the liquefaction pre-treatment, due to larger acid gas removal units and the requirement for a sulphur recovery step. The same economic forces also push the limits of required operating window to for example colder feed gas temperatures and higher operating pressures. The requirement for drying and mercury removal has remained basically constant.

2. Contaminants in Natural Gas Treating unit requirements are determined by the liquefaction unit requirements (water, CO2) , specifications of the LNG product (H2S, COS, organic sulphur compounds), material protection (mercury) and environmental restrictions( SO 2 and hydrocarbon emissions). In addition waste streams have also to fulfil minimum specifications. Typical LNG plant specifications are: Hydrogen Sulphide ex acid gas treating unit < 3.5 ppmv Carbon Dioxide ex acid gas treating unit < 50 ppmv Total Sulphur (Hydrogen Sulphide +Carbonyl Sulphide + Organic Sulphur Compounds ex acid gas treating unit) < 20 mg /Nm3 Total sulphur in fuel gas

< 300 ppmv (depending use)

SO 2 emission ex incinerator

< 250 mg /Nm3

Sulphur purity Sulphur recovery

> 99.9 %wt > 95 - 99.9 %

Water in LNG ex driers

< 0.5 ppmv

Mercury in LNG

< 0.01 µ g/Nm3

3. Liquefaction Pre-Treatment 3.1. Introduction Pre-treatment upstream a liquefaction unit traditionally consists of an acid gas removal step, in which CO2 and sulphur compounds (H2S, COS and mercaptans) are removed, a dehydration step and a mercury removal step. Where there are limitations on the SO 2 emissions, the removed sulphur components are recovered as elemental sulphur. Environmental limitations to hydrocarbon emissions can require incineration of CO2 acid gas even in the absence of sulphur compounds. Figure 1 shows a typical block diagram of a base load LNG train. The mercury removal step can be positioned upstream of the acid gas removal or downstream of the dehydration step.

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3.2. Process challenges In the recent years economic forces have led to gas fields with more challenging treating requirements to be considered for development. At the same time ever more emphasis has been put on identifying the lowest cost fit for purpose design package. As always, increasingly tighter environmental constraints, including Greenhouse emissions, have been applied to new projects. Most current plants treat gas with less than 10% CO2 but developments with up to 60% CO2 are now being considered. Fields containing higher levels of sulphur (H2S, COS, organic sulphur components) are being developed. Higher pressure operation and offshore treating are also being considered. In addition the combination of low feed temperatures and low CO2 content requires increased investment for most solvent based technologies. With the forces described above and the emergence of new treating technologies and vendors it becomes increasingly important to identify the best overall integrated solution for LNG pretreatment. An integrated process selection study should be carried out by a consultant with experience and operating knowledge of the various technologies and how they impact on and integrate with the overall LNG plant design.

3.3. Acid Gas Removal The wet absorption (solvent based) acid gas removal still remains clearly the most cost effective for base load LNG applications. Developments in cryogenic and membrane CO2 removal have yet to threaten the position of the solvent based processes when deep removal of CO2 for LNG production is required. Three basic types of liquid absorption processes are available: • Physical absorption processes, which use a solvent that physically absorbs CO2, H2S and organic sulphur components. Examples are the Purisol and Selexol processes. Physical solvents can be applied advantageously when the partial pressure of the contaminants are high, the treated gas specification is moderate and large gas volumes have to be purified. Physical solvents also absorb significant quantities of hydrocarbons, which obviously is a disadvantage. • Chemical absorption processes, which chemically absorb the H2S, CO2 and to some extent COS. Organic sulphur components do not chemically react with the solvent. Common examples are amine processes, using aqueous solutions of alkanol amines such as MEA, DEA, MDEA, DIPA and Flexsorb, and the carbonate processes, such as the Benfield process. Chemical solvents are specifically suitable when contaminants at relatively low partial pressure have to be removed to very low concentrations. Chemical solvents will not remove mercaptans down to low levels due to the low solubility of these components. An advantage however is that there is minimum co-absorption of hydrocarbons. Due to the chemical reaction between the solvent and CO2 and H2S, the regeneration energy requirements are normally higher than for a physical solvent. • Mixed solvents, are a mixture of a chemical and a physical solvent. The most widely known process is the Shell Sulfinol Process, which applies a mixture of sulfolane, water and DIPA (diisopropanolamine) or MDEA (methyldiethanolamine), Sulfinol-D and Sulfinol-M respectively. It combines to a large extent the advantages of a chemical with those of a physical solvent. One of the strengths of the Sulfinol process is the capability to simultaneously remove organic sulphur compounds and COS, which are not removed by pure chemical solvents. These characteristics make it the obvious choice for many natural gas treating problems. The Flexsorb SE process also combines sulfolane and an amine and is in many ways similar way to Sulfinol. In the oldest base load LNG plants MEA, DEA and carbonate processes are applied. In the early seventies the Sulfinol process was introduced and recently there has been one retrofit application of a formulated MDEA solvent (Badak, Indonesia).

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The Shell Sulfinol process is the most widely applied acid gas removal process in the base load LNG industry (~40% of current installed capacity) and has also been chosen for the majority of new projects currently under construction. In total there are about 200 Sulfinol units world wide. The main advantages of the process are: • the capability of simultaneously removing H2S, CO2, COS and organic sulphur components, • a high reliability, because sulfolane acts as a foam inhibitor resulting in a lower foaming tendency than aqueous amine solvents, and because of the low solvent corrosivity, • the low specific energy consumption, • the physical character of the solvent can be varied by choosing the relative amounts of water and sulfolane to trim the COS and mercaptans removal [9], • a lower water content in treated gas compared to aqueous amines • widely available solvent components not restricted to a proprietary supplier. • the ability to treat cold, low CO2 gas where rich solvent temperatures approach 20oC • the ability to operate at pressures well above 100 bara. A potential disadvantage is the somewhat higher hydrocarbon solubility. However, this solubility is still far less than that of a purely physical solvent and is normally considered acceptable. It is possible to remove H2S, CO2 and mercaptans using molecular sieves, however, this is not economical for large quantities of CO2 [1]. When H2S and mercaptans are removed with molecular sieves, the regeneration gas needs to be treated in a separate H2S and mercaptans removal unit.[17] This generally results in a more complex line-up and higher cost. The Sulfinol technology has recently been enhanced with a new patented technology for solvent reclamation. This process based on two stage wiped film evaporation provides a simple and efficient way to keep solvent in as new condition by separation of degradation products, heat stable salts and fine particulates with 95% recovery of the solvent components. F e e d G a s S u p p l y G a s R e c e i v i n g & M e t e r i n g

M e r c u r y R e m o v a l

A c i d G a s R e m o v a l F u e l

C O 2 S O 2

I n c i n e r a t i o n

S u l p h u r R e c o v e r y

g a s S u l p h u r

D i s p o s a l

D e h y d r a t i o n

U * * * *

tilities H T F Electricity W a t e r N i t r o g e n

M e r c u r y R e m o v a l P r e - T r e a t m e n t S e c t i o n H y d r o c a r b o n S e p a r a t i o n

F r a c t i o n a t i o n

L i q u e f a c t i o n

N G L

T r e a t m e n t

L N G

N G L

S t o r a g e

L N G

S t o r a g e

S h i p p i n g

N G L

P r o d u c t s

Figure 1. Typical base load LNG plant line-up.

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3.4. Sulphur Recovery If H2S is present in the feed, sulphur emissions will be subject to local legislative restrictions. The most appropriate process to convert the H2S and organic sulphur to elemental sulphur is mainly dependent on the quantity of sulphur present in the feed. Above about 15 t/d sulphur (figure 2), the common choice is the two stage Claus process, which typically achieves 95% sulphur recovery. To improve the overall sulphur recovery to 99.9 % a SCOT (Shell Claus Offgas Treating) unit [12] can be added. SCOT is one of the most widely applied Claus tail gas treating technologies with more than 180 units world-wide. A Claus unit produces significant quantities of medium and/or low pressure steam (~ 3 t steam/t of sulphur produced). CO2 removed from the natural gas and present in the acid gas acts as an inert gas in the Claus unit, increasing the size of the unit and reducing the sulphur recovery efficiency. If the feed to the Claus contains less than some 40 mol% H2S, there will be too many inerts present to maintain a stable flame in the main burner via a straight through Claus process. However, with special measures, e.g. enriched air, fuel gas co-firing, etc lower H2S concentrations can be accepted although efficiency may be reduced. To minimise Claus unit size and maximise efficiency the feed gas can be enriched by selectively removing the H2S from the regenerator off-gas in a second amine absorption step. The off-gas from the second amine regenerator will contain more than 40% H2S and is more suitable as Claus feed. For low H2S levels there are also other processes available, e.g. Selectox. For lower quantities of sulphur (< 5-15 t/d) the Claus unit can be replaced by a redox type process, such as the SulFerox process or the new biological Shell Thiopaq process, which are not restricted by a minimum H2S content in the feed gas. As stated above it is important to identify the best integrated and economic solution to the overall treating needs. The SCOT tail gas technology can readily be integrated with the main gas treatment, with or without an acid gas enrichment step and, if necessary, treatment of sour gas from molesieve regeneration.

1000000

15 ton su

H2S concentration (ppmv)

100000

Amine + Claus ( + SCOT)

lph Eff ur ect /d so f re ay c ove 5t r on y, t sul urn ph dow ur na /d nd ay CO

10000

2

SulFerox or kg sul Eff ph Shell ect ur s of /d wa 50 ay ter Thiopaq k sa 100

1000

gs

ulp

hu

r/

tur ati on

day

100

Disposable solids and liquids 10 0.01

0.1

Amine+SulFerox or Thiopaq 1

10

100

Gas Flow (MMNm3/day)

typical base load LNG plants

Figure 2. Process selection for H2S removal from natural gas

3.5 Integration possibilities with sulphur removal Various options exist to optimise the treating line-up in case of sulphur removal. In case of H2S only in quantities above, say, 15 tons/day (Figure 2), sulphur recovery is performed with a 2-stage Claus

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unit. When the recovery should be larger than 95 %, a tail gas treating unit is often required. Commonly applied processes are the SuperClaus and sub-dewpoint processes which are generally selected for efficiencies of around 99 %, while the SCOT technology is utilised for more than 99.5 % overall S-removal. When a SCOT process is selected, the absorption section of the SCOT can be easily integrated with the Sulfinol process in the acid gas section. In case the H2S/CO2 ratio in the feed gas is <0.2 mol/mol, the acid gas ex the acid gas removal plant will be difficult to process in a straight through Claus unit. A way to improve the Claus feed gas quality would be to remove CO2 from the acid gas before routing to the Claus unit. Two options to improve the situation exist: •

include an enrichment step in the acid gas removal scheme. An enrichment step includes a CO2 flash prior to solvent regeneration to reduce the amount of CO2 to the Claus unit. If required an H2S re-contactor is added to fulfil flash gas specifications.

•

enrich the acid gas before sending to the Claus unit in a dedicated system.

When an enrichment step is selected, this can be included in the acid gas removal step. A solvent which is selective for H2S is most effective. However, a selective solvent may not be the most appropriate choice for CO2 removal. The H2S/CO2 ratio and process conditions will determine the applicability of this option. Alternatively enrichment is achieved by a dedicated unit selectively absorbing H2S from the acid gas produced by the main CO2 removal unit. In this case a stand alone enrichment system is used. Due to the better selectivity this option may economically be more attractive [16]. When removing organic sulphur, partial mercaptan removal in a mixed solvent can be combined with removal of the rest of the mercaptan a separate molecular sieve bed (integrated in the dehydration unit). The molecular sieve regeneration gas is treated in a dedicated absorber column. This has been described in reference [17]. The regeneration gas may also be recycled to the main acid gas absorber. Optimisation of such a line up is very dependent on the level of organic sulphur in the feed gas and the total sulphur specification for the treated gas. Beside the optimisation in process line-up determined by component removal, optimisation based on energy requirements is possible. As example the use of the steam generated in the Claus unit may be sufficient to supply the regeneration heat to the dedicated tail gas unit. To determine the best line-up and process selection is very much dependent on the initial feed gas conditions, the treated gas specifications and environmental requirements. Before the treating cycle is selected, it is strongly recommended that an optimisation study is carried out to obtain the lowest capex/opex and largest operating window with respect to feed gas composition.

3.6 Incineration If the acid gas removal off-gas only contains CO2 it can be vented, however if H2S and/or aromatics are present, even in small amounts, the gas must be sent to an incinerator to prevent unsafe situations. The incinerator is normally of the thermal type, operating at about 800 oC, to achieve nearly complete H2S combustion (< 10 ppmv). If a Claus and/or SCOT unit are used the tail gas must also be incinerated. Modern adsorption materials also present the possibility of removing low levels of H2S by fixed beds where it is permitted to vent the balance of the acid gas stream. An interesting recent development is the option of using the hydrocarbon containing CO2 off-gas as fuel in steam boilers or furnaces. This concept can eliminate the need for an incinerator or allow for a considerably smaller incineration unit. With the increasing emphasis on reduction of greenhouse gas emission the option of acid gas recompression and subsequent re-injection is now being considered.

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3.7 Dehydration The sweet gas leaving the acid gas removal step is saturated with water because most sweetening solvents are aqueous solutions. First, the bulk of the water is condensed and separated from the gas stream by cooling. The cooling is limited to temperatures above the hydrate formation temperature (~20 oC). The water content is then further reduced to 0.5 ppmv by drying, normally with molecular sieves (4A-type). At least two of these dryers are required for each train; one is in adsorption service while the other is being regenerated by heated dry gas. The adsorption capacity of the molecular sieve decreases within two to four years to a level where it is necessary to replace the deactivated material. Recent developments have resulted in improved molecular sieves with higher absorption capacity and more resistance to degradation. The latest modelling techniques allow better optimisation of design and operation. Our research work has shown that deactivation is mainly due to a combination of two effects: (a) coke formation in the sieve and (b) caking due to hydrothermal instability of the sieve. Coke reduces the maximum water adsorption capacity, while caking leads to irreversible loss of drying capacity due to an increased residual water load after regeneration. With our specific knowledge on both the adsorption step and the regeneration step and the understanding of the deactivation mechanisms, we can optimise the operation of these units to be more cost effective. The combination of mercaptan removal and de-hydration in one step is discussed in section 3.5.

3.8 Mercury Removal Mercury removal is normally done with a fixed bed adsorption step. Commonly used adsorbents are sulphur impregnated carbons, in which the mercury reacts with sulphur to form the stable mercuric sulphide. A standard molecular sieve will also absorb Hg but regeneration is impossible. An alternative approach is the silver-impregnated molecular sieve (UOP HgSIV). In principle this molecular sieve can be regenerated, however the release of mercury from the molecular sieve bed would require dedicated material selections in the regeneration gas treating section. Integration of this Hg removal bed with the dehydration step is claimed to be possible [14]. Usually the preferred line-up has a simple dedicated mercury removal step (non-regenerative). The position of the mercury removal unit in the treating section depends on feed gas composition, the material from which it is constructed, and the water content of the gas. The mass transfer zone of the bed will be long at high water contents. In the area of mercury removal recent progress has mainly been in the detection and speciation of the different types of mercury. The current detection limit is 0.002-0.003 microgram/Nm3, which is an order of magnitude lower than some years ago. Generally mercury removal units remove the metallic mercury to below this detection limit.

4. Cost 4.1. Introduction The capital investment of a treating scheme is, of course, dependent on the type and concentrations of the contaminants in the feed gas and environmental targets. The costs of the dehydration and the mercury removal steps are mainly dependent on the gas throughput. The water content of the feed gas to the dehydration unit will not vary significantly if upstream cooling is applied to say about 20 oC. Dehydration can be significantly cheaper if a nonaqueous solvent can be used. The size of the mercury removal bed is mainly determined by contact time and hence throughput. The H2S, CO2 and mercaptan content of the gas will greatly influence the cost of the acid gas removal and the sulphur recovery units. Apart from the capital investment, the energy consumption,

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the operational cost of the unit and the hydrocarbon losses are important parameters. Integration of the processes can contribute considerably to the cost reductions.

4.2. Feed gas with only CO2 Table 1 gives an impression of the relative cost of the treating units for a typical base load LNG train (3 mtpa) as a function of the CO2 content and assuming negligible sulphur levels. The treating scheme is based on one Sulfinol-D CO2 removal train, molecular sieve dehydration and carbon bed mercury removal. The costs are expressed as a percentage of the cost of the combined processing facilities. Clearly the CO2 concentration in the feed gas has a major impact on the cost of the whole LNG train, whereas the cost of the dehydration and mercury removal unit remains fairly constant. The energy requirements obviously also increase with CO2 content. Where gas turbines are used for the liquefaction train and waste heat recovery (WHR) is applied, normally no or limited extra fuel gas is required over and above the gas turbine requirements. However, the cost of the WHR system increases with the energy demand. At high CO2 concentrations the equipment sizes may be close to the limit of what is technically possible to fabricate and install. Therefore a two train approach or a (Sulfinol-M) bulk CO2 removal step upstream of the Sulfinol-D deep removal step may be required. Table 2. Capital cost of the pre-treatment section relative to the total cost of the processing units for a typical LNG train (~3 mtpa LNG). 2 % CO 2 in feed 11 % CO 2 in feed 14 % CO 2 in feed Mercury Removal (%) 1.2 0.9 0.9 Dehydration (%) 4.5 3.5 3.5 C O 2 removal unit (%) 6.5 13.0 14.5

In previous work [1] the conclusion was drawn that for large gas volumes containing less than 1520 mol% CO2 the Sulfinol process was more economical than pure physical processes. Since 1996 several cases have been studied for the comparison of Sulfinol-D with alternative amine processes. The alternative processes comprise a representative modified MDEA solvent and an alternative mixed solvent. Results are shown below. Costs 13 % CO2 removal processes +/- 30% 120 Relative costs

100 80

Capex

60

Opex

40 20 0 Alt MDEA

Alt mixed solvent

SulfinolD

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Costs for 3 %v CO2 removal processes +/- 30 %

Relative costs

120 100 80

Capex Opex

60 40 20 0 Alt MDEA

Alt mixed solvent

SulfinolD

The formulated MDEA process figure is based on the most competitive MDEA based process as determined by recently completed process comparison studies. In the selection of the most attractive process, uncertainties in feed gas compositions (e.g. level of sulphur) and future requirements can often shape the ultimate choice.

4.3. Feed Gas with H 2S, CO2 and Organic Sulphur Components If, as in some recent projects, H2S and organic sulphur compounds are present in the feed gas to the LNG plant the cost of the gas pre-treatment can increase significantly. Where organic sulphur components are present, the Sulfinol process is often the preferred process as it combines H2S, CO2, COS and mercaptans removal in one step. This cannot be achieved with aqueous amines. In recent years this approach has been applied in the Shell Caroline plant [10], the Sexsmith gas plant [9] and the Qatargas LNG plant [13]. If the H2S and CO2 content of the feed gas is high, the mercaptans are removed without any additional cost. Only where the H2S and CO2 concentrations are low, will mercaptans removal determine the size of the acid gas removal unit. If there are restrictions on the SO 2 emissions, a sulphur recovery step is required. From figure 2 it can be concluded that above about 500 ppmv H2S in the feed gas, a Claus unit is almost the automatic choice. For a project producing 3 mtpa LNG, with 2 mol% CO2 and 2 mol% H2S in the feed, the sulphur production is about 400 t/d. In this case the cost of the Claus unit is similar to the cost of the acid gas removal unit. If the SO 2 emission limit cannot be met with just a Claus unit , a SCOT unit can be used to improve the sulphur recovery efficiency to 99.9%. This will obviously increase the overall cost. As a bonus the steam generated by the Claus unit may be sufficient to make the total treating scheme almost self-supporting with respect to energy consumption. As explained before in chapter 3, the cost of a Claus unit increases with decreasing H2S content of the acid gas feed. Figure 3 gives an impression of the capital cost of a sulphur recovery unit as function of the H2S content of the feed [2]. H2S enrichment of the Claus feed gas can be economical, especially if the enrichment is combined with a SCOT unit. The enrichment absorber and the SCOT absorber then share the same regenerator and solvent cascading is possible. This type of enrichment scheme will be applied in a new LNG project. Other enrichment schemes have been discussed before [2] and have been applied in e.g. the NAM Emmen Sulfinol-M plant [11]. It is very important for these applications that already in an early stage a good feed characterisation is made. The lowest foreseen H2S/CO2 ratio in the feed gas and the mercaptan content are critical to the final process selection [9] and developing the most cost effective design.

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1.9

Appr. relative capital investment per ton sulphur produced

1.8

1.7

1.6

1.5

1.4

1.3

1.2

1.1

1 0

10

20

30

40

50

60

70

80

90

100

% H2S in feed to Claus unit

Figure 3. Effect of H2S content of Claus feed on the relative cost of a Claus plant.

To illustrate the effect of H2S on the costs of a pre-treating section, the example represented in table 2 has been used with the addition of a 90 t/sd sulphur removal requirement. The adaptations in the line-up foresee in an enrichment section and a sulphur recovery unit. In view of expected environmental constraints a 99.9 % sulphur recovery has been taken in account. The results are represented in table 3. Table 3. Capital cost of the pre-treatment section relative to the total cost of the processing units for a typical LNG train (~3 mtpa LNG) with the addition of 90 t/d H2S removal capacity. 2 % CO 2 in feed 11 % CO 2 in feed 14 % CO 2 in feed Mercury Removal (%) 1.1 0.8 0.8 Dehydration (%) 4.3 3.3 3.3 H 2 S/CO 2 removal unit (%) 10.5 17.7 19.3

Comparing table 2 and 3 leads to the conclusion that the cost share for the gas treating step substantially increases with the addition of H2S removal capacity. Contributing factors are the need for enrichment (in this example) and the sulphur recovery. The effect of H2S removal on the gas treating processes is illustrated below for a typical aqueous MDEA (Adip) solvent and a mixed solvent. It should be realised that again the feed gas composition is of major importance. Costs for low CO2/1.3 %v H2S removal +/- 40 % 120

Relative costs

100 80 Capex

60

Opex

40 20 0 Alt mixed solvent

Aqueous MDEA

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In case the CO2 content is very high (in this example > 20 %v) a specific enrichment of the acid gas ex gas treating section is required to provide an acceptable Claus feed gas. Influence on the costs is represented in the next bar chart. Please note that this is a very severe example.

Relative costs

Costs for high CO2/1.3 %v H2S removal processes +/40 % with and without enrichment 200 180 160 140 120 100 80 60 40 20 0

Capex Opex

Alt mixed solvent

Aqueous MDEA

Alt mixed solvent + enrichment

Aqueous MDEA + enrichment

5. Developments in Pre-Treatment Membrane separation is based on thin barriers which allow preferential passage of components to separate multi-component feeds. The driving force for permeation is the difference in partial pressure between the feed gas and the permeate side of the membrane. Selectivity is achieved by differences is permeation rates. Membrane units for natural gas processing are commercially being used for CO2 / H2O removal to pipeline specification [15], production of CO2 for Enhanced Oil Recovery (EOR), unloading of existing treating plants, concentration of natural gas liquids. The use of membranes in large volume natural gas applications for e.g. CO2 removal is still restricted by the relatively high hydrocarbon losses and the fact that membranes do not have the same economy of scale as many other processes due to their modular nature. Current developments are therefore in improved permeance, reducing the required membrane area, and in improved CH4/CO2 selectivity to further reduce the hydrocarbon losses with the permeate. Inorganic membranes are an example of recent developments in this area [7]. Membranes can be economically used in the bulk CO2 removal step for natural gases containing high concentrations of CO2. Amine processes using partial flash regeneration, such as Sulfinol-M or MDEA, are the most economical solution for large volumes of natural gas containing less than about 15 mol% CO2 [6]. Above 15 mol%, a bulk CO2 removal step upstream of the amine unit can be attractive. Membranes can also be attractive for debottlenecking of existing CO2 removal plants, as long as the permeate can be used as fuel. Another area of interest is the use of membrane technology for LNG floaters, where movement can be significant. A disadvantage of conventional technology is that the internals, such as acid gas absorber trays and packing are sensitive to movements [8]. Several research institutes, such as GRI (USA) and TNO (Netherlands), are developing membrane alternatives for absorber trays, which are less sensitive to movement. An added advantage of the membrane contactor is the reduction of heavy hydrocarbon co-absorption, while the absence of direct gas-solvent contact removes the risk of foaming.

6. Technology Selection Simple rules for the selection of the most effective LNG pre-treatment technology are difficult to arrive at given the array of possibilities in feed gas conditions and resultant requirements for one or

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more of : one or two stage gas treatment, acid gas enrichment, mercaptan removal, sulphur recovery technology, Claus tail gas treatment, high pressure operation, low temperature operation etc. Added to this are other considerations such as reliability record, corrosion performance, quality of process guarantees and availability of professional after sales technical service. Given the magnitude of the investment in an LNG facility it is appropriate to carry out a rigorous treating process selection study to identify the most cost effective and fit for purpose treatment package. The consultant chosen to do the study must be experienced in the field of LNG gas pretreatment design, familiar with all the technologies available, able to discern effective integration options and ideally be able to draw upon actual operating experience to identify a robust operating line up. The investment in a high quality consultant to provide process selection will be well rewarded with a reliable, fit for purpose and cost effective treatment package.

7. Conclusion Up until recently the levels of CO2 and sulphur encountered in most LNG developments have allowed a fairly open selection of solvent based pre-treatment technology although preference has most often been given to well proven technologies. However with increased sulphur levels, mercaptans and more extreme operating requirements the available choices for a cost effective integrated treating package narrow dramatically. Use of the Shell Sulfinol technology can be the decidedly more attractive option in these cases. It is critical that each new gas development undergo a rigorous selection study to identify the most cost effective and fit for purpose treating package.

8. References [1] Klein, J.P., Verloop, J., The purification of Natural Gas prior to liquefaction, Paper presented at the Second Iranian Congress of Chemical Engineering, Tehran, 11th-14th May 1975 [2] McEwan, M.C., Marmin, A., Developments in LNG Treating, Proceedings of LNG-6 Kyoto April 1980 Vol.II Appendix Paper 3. [3] Peebles, M.W.H., Natural Gas Fundamentals, Shell International Gas Limited, 1992 [4] Avidan, A.A., et al., LNG links remote supplies and markets, Oil & Gas Journal, June 1997,pp 54-59 [5] Bellow, E.J., et al., Technology advances keeping LNG cost-competitive, Oil & Gas Journal, June 1997, pp 74-78. [6] McKee, R.I., et al., CO2 removal: membrane plus amine, Hydrocarbon Processes, April 1991, pp 63-65 [7] Minutes Gas Utilisation Research Forum, June 1997, Copenhagen, Denmark. [8] Naklie, M..M., Design advanced for large-scale economic, floating LNG plant, Oil & Gas Journal, June 1997, pp 66-69 [9] Johnston, A.L., Moore, B.K., Balancing cost with process, resident and regulatory needs, Laurence Reid Gas Conditioning Conference paper, March 1997. [10] Cheng, N. et al., Shell Canada’s Caroline Gas Plant reaches smooth operation, Laurence Reid Gas Conditioning Conference paper, March 1997 [11] Taylor, N.A., et al., Gas-desulfurisation plant handles wide range of sour gas compositions, Oil & Gas Journal, August 1991, pp 57-59 [12] Wetzels, M.L.J.A., et al., The SCOT Process: still very much alive nearly twenty years after its first introduction, Western Research Seminar, November 1990, Budapest. [13] Oil & Gas Special: Natural gas in the Middle East, February 1997, pp 49-64 [14] Holmes, E.S. et al., Process Improvements in acid gas and mercury removal, AIChE 1994 Spring National Meeting, Atlanta, April 1994 [15] Cook, P.J., Losin, M.S., Membranes provide cost-effective natural gas processing, Hydrocarbon Processing, April 1995, pp 79-84 [16] Abdul Redha Abdul Rahman, Qatar Liquefied Gas Co., Doha, Qatar & Masayuki Ishikura, Chiyoda Corp., Yokohama, Japan, 1’LNG from giant North field: Qatar Gas LNG, 1998 presented LNG 12 conference, Perth, May 1998 [17] Clarke D.S., Sibal P.W., Gas Treating Alternatives for LNG Plants, GPA, March 16-18, 1998, Dallas Texas

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